Hydrocracking process with pre-hydrofining



Jan. 16, 1958 5.1. YOUNG ET Al- 3,35433 HYDROCRACKING PROCESS WITH PRE-HYDROFINING Filed Sept. 25, 1964 United States Patent O 3,364,133 HYDROCRACKING PROCESS WITH PRE-HYDROFINING Billy J. Young and John H. Duir, Fullerton, Calif., assignors to Union Oil Company of California, Los

Angeles, Calif., a corporation of California Filed Sept. 23, 1964, Ser. No. 398,521 Claims. (Cl. 208-89) This invention relates to the catalytic hydrocracking of high-boiling hydrocarbons to produce relatively lower boiling hydrocarbons, boiling for example in the gasoline and/or jet fuel range. More specifically, the invention is directed to a novel two-stage prehydrotining treatment, wherein the second hydroning stage is operated integrally with a succeeding hydrocracking stage. An integral hydroiining-hydrocracking sequence, as the term is employed herein means that ei-uent from the hydroining zone is sent directly to the hydrocracking zone without signilicant interstage treatment or purification, so that the hydrocracking is carried out in the presence of ammonia formed in the preceding hydroiining stage.

Brieliy stated, the invention comprises the following steps:

1) Subjecting a nitrogen-containing hydrocarbon feedstock t0 catalytic hydrolinin-g in a separate hydrotining unit, preferably operated at rel-atively low pressures and high space velocities, to effect a limited hydrofining involving the conversion to ammonia of only about 3090% of the initial feed nitrogen content;

(2) Subjecting effluent from the first hydrofining treatment to a purification procedure for removing at least the major portion of ammonia contained therein;

(3) Subjecting the dea mrnoniated eliiuent from step (2) to a second hydroiining operation, preferably at relatively high pressures, and under conditions adjusted to effect a further substantial reduction in organic nitrogen content, preferably to less than about 50 parts per million; and

(4) Subjecting etliuent from second hydroiining step, together with the ammonia formed therein, to catalytic hydrocracking at high pressures, preferably the same as in the second-stage hydro-finer, under conditions adjusted to provide a substantial conversion per pass to desired low-boiling hydrocarbons.

The overall object of the invention is to provide a hydrcining-hydrocracking combination process wherein the size of the high-pressure reactors is reduced to a minimum, while at the same time obtaining the economic advantages of an integral hydrofning-hydrocracking combination. A specific object is to provide economical means for hydrocracking feedstocks containing relatively large amounts of organic nitrogen, i.e., above about 0.30 weight percent. Another object is to increase the octane number of the gasoline product resulting from an integral hydroning-hydrocracking combination. Still another object is to provide convenient means for controlling the ammonia and nitrogen content of feed to the hydrocracking zone in an integral hydrotining-hydrocracking combination. Other objects will be apparent from the more detailed description which follows.

It is known in catalytic hydrocracking that the nitrogen compounds contained in hydrocarbon feedstocks exert a marked poisoning elfect upon hydrocracking catalysts. To avoid this problem, it has become standard practice 3,354,133 Patented Jan. 16, 1958 ICC to subject nitrogen-containing feedstocks to a preliminary catalytic hydrofining treatment so as to reduce the nitrogen content to below about 10 parts per million. Typically, this is accomplished in a non-integral hydrofiner, wherein the hydroliner etiiuent is cooled, condensed, depressured and water-washed to remove ammonia, then re-pressured and re-heated for the hydrocracking stage. These interstage treatments, particularly the cooling and re-heating involved, add greatly to the expense of the overall operation.

In copending application, Ser. No. 142,182, now U.S. Patent No. 3,159,568, filed Oct. 2, 1961, it is proposed to avoid this expense by eliminating interstage treatment, and passing the entire hydroiiner etiluent through the succeeding hydrocracker. This proposal was based upon the discovery that the ammonia formed during hydroiining is much less deleterious than organic nitrogen compounds in their effects upon hydrocracking catalyst activity. Specifically, it was found that the effects of ammonia are reversible, and moreover can be much more readily overcome by increasing hydrocracking temperat-ures than can the effects of org-anic nitrogen compounds. Organic nitrogen compounds tend to bring about progressive catalyst deactivation until temperatures in the range of about 850 to 900 F, are reached. Temperatures in this range are undesirable because of the increased catalyst coking rates and production of light hydrocarbons in the C1-C4 range. When the organic nitrogen compounds were converted to ammonia, hydrocracking could successfully be carried out at temperatures in the range of about to 780 F. without encountering progressive deactivation resulting from nitrogen poisoning.

It has now been found, however, that in cases where the ammonia partial pressure in the hydrocracking zone exceeds about 4-5 p.s.i., it is diicult to maintain desired conversion levels at temperatures below about 780 F. Optimum hydrocracking temperatures are considered to lie in the range between about 680 and 780 F., for in this range product quality (octane value) is optimum when correlated with yield structure. At hydrocracking temperatures below about 680 F., there is a marked lowering in octane values; at temperatures above about 780 F. the increased rate of production of coke and light hydrocarbon gases offsets the further slight increase in octane value obtained. It would therefore be desirable to maintain only suiiicient ammonia in the hydrocrackin-g zone to permit the desired conversion per pass to gasoline at temperatures in the range of about 680 to 780 F. This invention provides economical means for achieving this objective.

The organic nitrogen compounds contained in miner-al oil feedstocks are varied in nature. The less refractory compounds are easily decomposed during hydrotining at relatively low pressures. Other more refractory nitrogen compounds, particularly those found in the heavier portions of the feedstock, are very resistant to decomposition at low pressures, 'out can be effectively decomposed at higher hydroiining pressures. The cost of high-pressure hydroning, however, increases markedly as the size of the reactor increases. According to the present invention, high-pressure hydrotining is utilized only for completing the decomposition of refractory nitrogen compounds after most of the less refractory nitrogen compounds have been decomposed in a more economical low-pressure reactor. This permits a substantial increase in space velocities, and

3 thus minimizes catalyst and reactor volume required in the high-pressure hydroner.

At the same time, the degree of denitrogenation which takes place in the low-pressure hydroiner can be independently controlled, so as to leave in the effluent therefrom a controlled amount of organic nitrogen calculated to provide, through its decomposition in the high-pressure liydroner, the ammonia partial pressure in the hydrocracking zone (normally about 0.5-5 p.s.i.) which is desired for maintaining an optimum hydrocracking temperature therein. Thus, by gradually increasing the severity of hydroning in the low-pressure reactor during a run, it is possible to gradually decrease the resultant partial pressure of ammonia in the hydrocracker during the run so as to compensate for catalyst deactivation while maintaining a substantially constant desired ternperature in the hydrocracker throughout the run. The end result is a more uniform product quality from the hydrocracker at an optimum yield-octane structure.

For a more detailed description of the invention reference is made to the accompanying drawing which is a ow `sheet illustrating a preferred modification thereof. The initial feedstock is brought in through line 2, mixed with recycle and fresh hydrogen from line 4, preheated to incipient hydrofining temperatures in preheater 6, and passed into the low-pressure hydroner 8 containing a bed of granular hydroning catalyst 10. Suitable hydrofning conditions in hydroner 8 are as follows:

LOW-PRESSURE HYDROFINING CONDITIONS The above conditions should be suitably correlated so as to effect conversion of at least about 30% and up to about 90% of the organic nitrogen to ammonia, and in any case to reduce the organic nitrogen content to below about 0.30 weight percent. Suitable hydroning catalysts are described hereinafter.

The effluent from hydroner 8 is withdrawn through line 12 and condensed in cooler 14. The condensed efuent in line 16 is then subjected to water-washing by means of water injected via line 18, and the mixture is then transferred to high-pressure separator 20 from lwhich hydrogen-rich recycle gas is withdrawn via line 22, and mixed with fresh make-up hydrogen from line 24 for recycle via line 4 as previously described. Spent wash water containing dissolved ammonia, and some of the hydrogen sulde formed in hydroner 8 is withdrawn via line 26. The high-pressure liquid hydrocarbon condensate in separator 20 may either be passed directly to the highpressure hydroiiner 28 via lmes 30 and 32, or it may be depressured via line 34 into low-pressure separator 36 from which light hydrocarbon flash gases are exhausted via line 38. Ordinarily it is not necessary to utilize the low-pressure separator 36, but it may be desirable in some cases, as where hydroiining conditions in hydrofner 8 are relatively severe resulting in the formation in large amounts of Cl-Ca hydrocarbons. These light hydrocarbons, if present in excessive amounts, tend to reduce the efficiency of hydrotining in high-pressure hydroner 28 due to lowering of the hydrogen partial pressure therein. If low-pressure Separator 36 is employed, the low-pressure condensate therein is withdrawn via line 40 and passed to high-pressure hydroiiner 28 via line 32 as previously described. In either case, a feed transfer pump 42 is provided in order to maintain the desired high pressure in hydroner 28.

The ammonia-free condensate in line 32 is blended with fresh and recycle hydrogen from line 44, and the mixture is then passed into high-pressure hydroner 28 via preheater 46. Hydroner 28 also contains a granular bed of hydroner catalyst 48, which may be the same as or different thanthe catalyst employed in hydroner 8. Operating conditions in hydrotner 48 fall within the following ranges:

HIGH-PRESSURE HYDROFINING CONDITIONS The above conditions are suitably correlated so as to reduce the organic nitrogen content of the feed to below about 50, and preferably below about 30 parts per million. In cases where the hydroning severity in -low-pressure hydroner 8 is gradually increased during the run, so as to provide a gradually decreasing ammonia concentration in the hydrocracker to compensate for catalyst deactivation therein, there will be a concomitant decrease in organic nitrogen content of the eiiiuent from hydrofner 28 during the run, down to below about 10 parts per million or less at end-of-run conditions.

Eluent from hydroner 48 is withdrawn via line 50, and may be transferred in toto to hydrocracker 52 via lines 54 and 56 (valve 58 open, valve 60 closed). In some cases the hydroner eluent in line 50 may comprise a liquid phase, especially where the original feedstock is relatively high boiling. Where such a liquid phase is present, it will include most of the heavy ends of the feed material, and it is this heavier portion which is most difcult to denitrogenate. In such cases, instead of transferring the entire effluent to hydrocracker 52 via line 54, iti may be more desirable to adopt an alternate procedure by closing valve 58 and opening valve 60, thus diverting.

the effluent via line 62 to a high-pressure hot separator K 64, from which the hot vapor-phase eilluent is withdrawn via line 66 and sent to hydrocracker 52 as previously described. The liquid phase product in separator 64 may be withdrawn via line 68, and all or a portion thereof recycled via line 70 to high-pressure hydroiiner 28. By this procedure the heavy portion of the feedstock is hydrofined at a lower effective space velocity than the vapor phase portion, and by this means the light and heavy portions of feeds may be denitrogenated to more nearly equal degrees. Any portion of the liquid phase effluent which is not recycled via line 70 may be transferred directly via lines 72 and 56 to hydrocracker 52.

The total lhydrofiner eiliuent in line S6 is blended with recycle hydrogen from line 74 (if desired for temperature control and/or make-up purposes), and with recycled unconverted oil from line 76, and lthe resulting mixture is then passed via preheater 78 into hydrocracker 52, containing a bed of granular hydrocracking catalyst 80. Operr ating conditions in hydrocracker 52 fall within the following general ranges:

HYDROCRACKING CONDITIONS Operative Preferred Average Bed Temp., F 600-850 S80-780 Pressure, p.s.i.g 8003, 500 1, 000-2, 5001 LH 0. 5-5 l-4 Hg/Oil Ratio, M S.c.f./b- 1-20 1-12 Effluent from hydrocracker 52 is withdrawn via line 82 and condensed in cooler 84. 'I'he resulting two-phase mixture may then be washed with water injected via line 86, and transferred to high-pressure separator 88, from which substantially ammonia-free recycle gas is withdrawn via line 44 for recycle to hydroner 28 as previously described. Spent Wash-water containing ammonia and small amounts of hydrogen sulde is withdrawn via line 92.

High-pressure condensate in separator 88 is depressured `via line 94 into low-pressure separator 96, from which light hydrocarbon gases may be exhausted via line 98. Low-pressure condensate in separator 96 is then transferred to distillation column 100 via line 102. The final gasoline product, or other desired low-boiling hydrocarbon product, is withdrawn as overhead from column 100 via line 104, and un-converted oil is withdrawn via line 76 for recycle to hydrocracker 52 as previously described. As an alternate, the bottoms product from column 100 may be withdrawn and utilized as diesel or jet fuel if desired, or transferred to a second hydrocracker for further conversion to low-boiling hydrocarbons in the absence of ammonia and at relatively lower temperatures than were employed in hydrocracker 52.

The hydroning catalyst to be used in the above hydroning units may comprise any of the oxides and/or sulides of the transitional metals, and especially an oxide or sulde of a Group VIII metal (particularly iron, cobalt or nickel) mixed with an oxide or sulfide of a Group VI-B metal (preferably molybdenum or tungsten). Such catalysts preferably are supported on an adsorbent carrier in proportions ranging between about 2% and 25% by weight. Suitable carriers include in general the difficultly reducible inorganic oxides, e.g., alumina, silica, zirconia, titania, clays such as bauxite, bentonite, etc. Preferably the carrrier should display little or no cracking activity, and hence highly acidic carriers having a Cat-A cracking activity index above about 20 are to be avoided. The preferred carrier is activated alumina, and especially activated alumina containing about 3-15% by weight of coprecipitated silica gel.

The preferred hydroning catalyst consists of cobalt sulde or oxide plus molybdenum sulfide or oxide supported on silica-stabilized alumina. Compositions containing between about 1% and 5% of Co, 3% and 20% of Mo, 3% and 15% of SiOg, and the balance A1203, and wherein the atomic ratio of Co/Mo is between about 0.2 and 4, are specifically contemplated.

The hydrocracking catalyst to be employed herein may consist of any desired combination of a refractory cracking base with a suitable hydrogenating component. Suitable cracking bases are those having a Cat-A activity index above about 20, and preferably above 30, including for example mixtures of two or more diicultly reducible oxides such as silica-alumina, silica-magnesia, silica-zirconia, alumina-boria, silica-titania, silica-zirconia-titania, acid-treated clays and the like. Acidic metal phosphates such as aluminum phosphate may also be used. The preferred cracking bases comprise partially dehydrated, zeolitic, crystalline molecular sieves having relatively uniform pore diameters of about 6-14 A., and comprising silica, alumina and one or more exchangeable zeolitic cations. These crystalline zeolites may be used as the sole cracking base, or they may be mixed with one or more of the amorphous cracking bases such as silicaalumina cogel.

A particularly active and useful class of molecular sieve cracking bases are those having a relatively high Si02/ A1203 mole ratio, e.g., between about 3 and 10. Suitable zeolites found in nature include for example mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite, and faujasite. Suitable synthetic molecular sieve zeolites include for example those of the B, X, Y, and L crystal types, or syntheic forms of the natural zeolites noted 6 above, especially synthetic mordenite, The preferred zeolites are those having crystal pore diameters between about 8-12 A., wherein the SiO2/AI2O3 mole ratio is about 3-6. A prime example of a zeolite falling in this preferred group is the synthetic Y molecular sieve.

The naturally occurring molecular sieve zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic molecular sieves normally are prepared in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ion-exchanged out with a divalent metal, or with an ammonium salt followed by heating to decompose the zeolitic ammonium ions, leaving in their place hydrogen ions and/ or exchange sites which have actually been decationized by further removal of water:

Mixed divalent metal-hydrogen zeolites may be prepared by ion exchanging first with an ammonium salt, then partially back exchanging with a divalent metal salt, and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen forms can be prepared by direct acid treatment of the alkali metal sieves. Hydrogen or decationized Y sieve zeolites of this nature are more particularly described in Belgian Patents Nos. 598,582, 598,682, 598,683 and 598,686, and U.S. Patent No. 3,130,- 006.

There is some uncertainty as to whether the heating of the ammonium zeolites produces a hydrogen zeolite or a truly decationized zeolite, but it is clear that, (a) hydrogen zeolites are formed upon initial thermal decomposition of the ammonium zeolite, and (b) if true decationization does occur upon further heating of the hydrogen zeolites, the decationized zeolites also possess desirable catalytic activity. Both of these forms, and the mixed forms, are designated herein as being metal-cationdeficient.

The foregoing cracking bases are compounded, as by impregnation, with from about 0.5% to 25% (based on free metal) of a Group VI-B and/ or Group VIII hydrogenating metal promoter, e.g., an oxide or sulfide of chromium, tungsten, cobalt, nickel, or the corresponding free metals, or any combination thereof. Alternatively, even smaller proportions, between about 0.05% and 2% of the metals platinum, palladium, rhodium or iridium may be employed. The oxides and sulfdes of other transitional metals may also be used, but to less advantage than the foregoing.

In the case of zeolitic type cracking bases, it is desirable to deposit the hydrogenating metal thereon by ion exchange. This can be accomplished by digesting the zeolite with an aqueous solution of a suitable compound of the desired metal, wherein the metal is present in a cationic form, and then reducing to form the free metal, as described for example in Belgian Patent No. 598,686.

Although as indicated above, substantially any hydrocracking catalyst may be used herein, it is not to be as- Sumed that all such catalysts are equivalent, or that they will all give a commercially feasible process. For hydrocracking at below about 850 F. as required herein, and in the presence of ammonia, it is highly desirable to use catalysts which have both a high cracking activity and high hydrogenating activity; otherwise it will be necessary to employ uneconomical low space velocities. For economical processes operated at above about 0.5 LHSV, the preferred catalysts are composed of a Group VIII noble metal, e.g., platinum, palladium, rhodium, iridium or ruthenium, combined by ion exchange with one of the zeolitic molecular sieve cracking bases of the Y crystal type wherein the zeoliticcations are predominantly hydrogen and/or a divalent metal such as calcium, magnesium, or zinc. The more conventional catalysts such as platinum on silica-alumina gel, or nickel on silica-alumina gel, will require low space velocities, in general below about 0.5 in order to achieve the desired conversion at the specified hydrocracking temperatures.

The feedstocks which may be treated herein include in general any mineral oil fraction having an initial-boilingpoint above the end-boiling-point of the desired product, and having an end-boiling-point up to about 1,000 F. This includes straight run gas oils, coker distillate gas oils, deasphalted crude oils, cycle oils derived from catalytic or thermal cracking operations and the like. These fractions may be derived from petroleum crude oils, shale oils, tar sand oils, coal hydrogenation products and the like. Specilically, it is preferred to employ feedstocks boiling between about 400 and 900 F., having an API gravity of 20 to 35, and containing at least about 20% by volume of acid-soluble components (aromatics-i-oletins). Such oils may also contain from about 0.1% to of sulfur and from about 0.01% to 2% by weight of nitrogen.

The following example is cited to illustrate the invention Vand the results obtainable, but is not to be construed as limiting in scope.

Example This example illustrates process conditions and results obtainable in an extended hydrocracking run carried out as described in connection with the drawing, with lowpressure hydrofiner eti-'tuent being transferred directly from separator 20 to the high-pressure hydroning unit, and with total eiuent from the high-pressure hydroner being transferredA directly to the hydrocracker. The initial feedstock is a blend of straight-run gas oil and catalytic cracking cycle oil, having an API gravity of about 23.2, a boiling range of about 420 to 850 F., and containing about 1.06 weight-percent sulfur and 0.9 weight-percent nitrogen. The hydroning catalyst employed in each of the hydroners is the sulded equivalent of 3% cobalt oxide and molybdenum oxide supported on an alumina carrier stabilized by the addition of about 5% SiO2, in the form of a pellets. The hydrocracking catalyst is a co-pelleted mixture of about 80 weight-percent of a Y molecular sieve zeolite containing 0.5 Weight-percent palladium, and weight-percent of an activated alumina binder containing 0.3 weight-percent palladium. The Y molecular sieve cracking base has a SO2/A1203 mole ratio of about 4.7, about 35% of the zeolitic ion exchange capacity thereof being satisfied by magnesium ions (3 weight-percent MgO), about 10% by sodium ions,

and the remainder by hydrogen ions. Process conditionsY are adjusted so as to reduce the organic nitrogen content of the feed to about 0.3 weight-percent in the low-pressure hydroner, and to about 20 parts per million in the highpressure hydrotner. Hydrocracking conditions are adjusted to provide about 60 volume-percent conversion per pass. To achieve these objectives, the major process conditions in the various units are as follows:

PROCESS CONDITIONS At end-of-run conditions, product distribution and quality is approximately as follows:

8 Dry gas make (C1-C3), s.c.f./b. freshV feed 120 Liquid yields, vol. percent fresh feed:

Butanes 19 Pentanes 15 C6 16 Cq-400 F. gasoline 72 C4-400 F. gasoline 122 C5+gasoline octane No., F-l-l-3 ml. TEL 91 Y The relatively high octane'value of the gasoline product is attributable in part to the use of low pressures during 15 the first hydrolining zone. If all of the hydroning were `performed at the -high pressure of 1,500 p.s.i.g., la gasoline product about two octane numbers lowerwould be obtained.

It is not intended that the invention should be limited 20 to the details described herein, since many variations may be made by those skilled in the art without departing from the scope or spirit of the following claims.

We claim: Y 1. A process for hydrocracking a mineral oil feedstock in the optimum hydrocracking temperature range of about 680-780. F. to obtain a relatively high-octane gasoline with a minimum production of light hydrocarbon gases,

said feedstock boiling above the gasoline range and containing at least about 0.3% by weight of organic nitrogen, 0 which comprises:

(1) contacting said feedstock plus added hydrogenV with a hydrofim'ng catalyst at elevated temperatures and pressures in a lirst hydroning zone and controlling the severity of hydroiining so as to (a) effect about 3090% denitrogenation of said feedstock, and

(b) leave in said feedstock a remaining amount of organic nitrogen calculated to yield in step (3) an amount of ammonia sufficient to control the hydrocracking temperature in step (4) within the limits of 40 about 680-780 F., said remaining amount of organic tures in a second hydroning `zone'maintained at a pressure substantially higher than the pressure in said lirst hydrofning zone, to reduce the organic nitrogen content thereof to below about 50 p.p.m. with resultant formation of ammonia;

(4) passing eflluent from said second hydrofining zone,

including ammonia formed therein and hydrogen through a bed of hydrocracking catalyst at an elevated pressure, temperatures between about 680 and 780 F. and a space velocity between about 0.5 and 5, said conditions being correlated to give about 30- volume-percent conversion per pass to gasoline; and (5) recovering high-octane gasoline from said hydrocracking step. 5 2. A process as defined in claim 1 wherein the hydroning catalyst employed in each of said hydroining zones is a composite of a Group VI-B metal oxide and/or sulfide with a Group VIII metal oxide and/or sulfide supported on an adsorbent carrier having a Cat-A crack- 70 ing activity index below about 20, and wherein said hydrocracking catalyst comprises a Group VIII metal hydrogenating component supported on a cracking base having a Cat-A cracking activity index above about 20.

3. A process as defined in claim 2 wherein said hydrocracking catalyst comprises a Group VIH noble metal 9 hydrogenating component supported on a zeolitic aluminosilicate molecular sieve cracking base having a SO2/A1203 moleratio between about 3 and 10, and comprising zeolitic cations selected from the class consisting of hydrogen ions and divalent metal ions.

4. A process as dened in claim 1 wherein said hydrocracking step (4) is continued for a period of at least about 3 months while maintaining substantially constant conversion per pass therein by incrementally raising the hydrocracking temperature within the stated range of 680-780 F. to compensate for catalyst deactivation.

5. A process as defined in claim 1 wherein said hydrocracking step (4) is continued for a period of at least about 3 months while maintaining a substantially constant hydrocracking temperature within the stated range of 680-780 F., and wherein hydrocracking catalyst de- References Cited UNlTED STATES PATENTS 3,256,178 6/1966 Hass et al. 208-89 2,911,356 11/1959 Hanson 208-110 3,132,089 5/1964 Hass et al. 208-89 SAMUEL P. JONES, Primary Examiner.

DELBERT E. GANTZ, Examiner 

1. A PROCESS FOR HYDROCRACKING A MINERAL OIL REEDSTOCK IN THE OPTIMUM HYDROCRACKING TEMPERATURE RANGE OF ABOUT 680-780*F. TO OBTAIN A RELATIVELY HIGH-OCTANE GASOLINE WITH A MINIMUM PRODUCTION OF LIGHT HYDROCARBON GASES, SAID FEEDSTOCK BOILING ABOVE THE GASOLINE RANGE AND CONTAINING AT LEAST ABOUT 0.3% BY WEIGHT OF ORGANIC NITROGEN, WHICH COMPRISES: (1) CONTACTING SAID FEEDSTOCK PLUS ADDED HYDROGEN WITH A HYDROFINING CATALYST AT ELEVATED TEMPERATURES AND PRESSURES IN A FIRST HYDROFINING ZONE AND CONTROLLING THE SEVERITY OF HYDROFINING SO AS TO (A) EFFECT ABOUT 30-90% DENITROGENATION OF SAID FEEDSTOCK, AND (B) LEAVE IN SAID FEEDSTOCK A REMAINING AMOUNT OF ORGANIC NITROGEN CALCULATED TO YIELD IN STEP (3) AN AMOUNT OF AMMONIA SUFFICIENT TO CONTROL THE HYDROCRACKING TEMPERATURE IN STEP (4) WITHIN THE LIMITS OF ABOUT 680-780*F., SAID REMAINING AMOUNT OF ORGANIC NITROGEN BEING BETWEEN ABOUT 0.03 AND 0.3 WEIGHT PERCENT; (2) TREATING THE EFFLUENT FROM SAID FIRST HYDROFINING ZONE TO SEPARATE AMMONIA THEREFROM AND TO RECOVER A PARTIALLY REFINED OIL; (3) CONTACTING SAID PARTIALLY REFINED OIL PLUS ADDED HYDROGEN WITH A HYDROFINING CATALYST COMPRISING A GROUP VI-B METAL OXIDE AND/OR SULFIDE PLUS A GROUP VIII METAL OXIDE AND/OR SULFIDE AT ELEVATED TEMPERA- 